Hybrid high-temperature swing adsorption and fuel cell

ABSTRACT

Systems and methods are provided for combined cycle power generation while reducing or mitigating emissions during power generation. Recycled exhaust gas from a molten carbonate fuel cell power generation reaction can be separated by using a swing adsorption process so as to generate a high purity CO 2  stream while reducing or minimizing the energy required for the separation and without having to reduce the temperature of the exhaust gas. A high temperature adsorption reactor adsorbs the CO 2  and recovers H 2  from an exhaust gas of a first molten carbonate fuel cell at a high temperature and at a low pressure. The reactor passes along the adsorbed CO 2  to a cathode and the recovered H 2  to an anode of a second molten carbonate fuel cell for further power generation. This can allow for improved energy recovery while also generating high purity streams of CO 2  and H 2 .

CROSS-REFERENCE TO RELATED APPLICATIONS

This application claims priority to U.S. Provisional Application Ser. No. 62/256,330 filed on Nov. 17, 2015, herein incorporated by reference in its entirety.

FIELD

In various aspects, the invention is related to a pressure swing adsorption (PSA) application in a system for generating power by fuel cell technology. Specifically, the invention is related to an application of a low pressure high temperature PSA reactor which adsorbs gases that result from power generating reactions that occur in fuel cells.

BACKGROUND

Capture of gases emitted from power plants that include fuel cell generated power is an area of increasing interest. Power plants based on reactions that occur within molten carbonate fuel cells generate carbon dioxide (CO₂) as a by-product of the reaction. Historically, CO₂ has been either released into the atmosphere after the reaction takes places or prepared for storage by expensive and energy demanding processing prior to the storage. Therefore, traditional disposal of CO₂ in the process of power generation has been energy consuming. As a result, it is becoming increasingly desirable to identify alternative ways to process and recycle CO₂ generated during the fuel cell reactions.

Moreover, certain types of fuel cell technologies require CO₂ as one of the reactants in the process of power generation. At the same time, the emission gases that include CO₂ normally do not produce separated and isolated CO₂ ready to be used in a chemical reaction required to generate power. This poses an additional challenge where, while recycling CO₂ into molten carbonate fuel cells as a reactant could decrease the need for storing it, the process of separating CO₂ out of the emission mixture to be usable as a reactant draws a significant amount of energy. Thus, a need arises to develop new ways to inexpensively and efficiently prepare CO₂ to be fed back into the power generation process as a reactant.

Further, in molten carbonate fuel cells, hydrogen (H₂), the fuel gas, often does not react entirely, but is released as an emission gas. In order to maximize utilization of H₂ in the power plant, unreacted H₂ would need to be purified and fed back into the fuel cells.

One of the common problems of the task of preparing CO₂ for storage and the task of separating the reactant gases in order to convert them into suitable reactants is that, because they are portions of the emission mixture, they are available at high temperatures. Further, in instances where this preparation and separation is performed by pressure swing adsorption (PSA), for example, the emission gas at the outlet of a molten carbonate fuel cell flows at a relatively low pressure (˜1 bara). Thus, in view of the commonality of these processes, which are typically integrated in a single power generation system, there is a need for a solution that would address the discussed problems while minimizing the footprint of the system and while increasing its energy efficiency.

SUMMARY

Systems and methods are provided for combined cycle power generation while reducing or mitigating emissions during power generation. Recycled exhaust gas from a molten carbonate fuel cell power generation reaction can be separated by using a swing adsorption process so as to generate a high purity CO₂ stream while reducing/minimizing the energy required for the separation and without having to reduce the temperature of the exhaust gas. This can allow for improved energy recovery while also generating high purity streams of carbon dioxide and hydrogen.

In one embodiment of the present invention, a method for optimizing emission reduction of a power generation system, comprises: passing hydrocarbons to a reformer for processing and producing H₂, passing the produced H₂ to a first anode of a first molten carbonate fuel cell, passing CO₂ and O₂ to a first cathode of the first molten carbonate fuel cell for processing and producing carbonate ions, transporting the produced carbonate ions from the first cathode to the first anode through a first electrolyte of the first molten carbonate fuel cell, generating power and an exhaust gas by reacting the H₂ with the transported carbonate ions in the first anode of the first molten carbonate fuel cell, the exhaust gas including CO₂ at a temperature greater than 600° C., passing the exhaust gas of the of the first molten carbonate fuel cell to a first swing adsorption reactor, and adsorbing the CO₂ from the exhaust gas on a first adsorbent material of the first swing adsorption reactor.

The method for optimizing emission reduction of a power generation system may further comprise purging the first swing adsorption reactor with steam and removing the adsorbed CO₂ and/or recovering a H₂ stream unadsorbed by the first swing adsorption reactor. The first swing adsorption reactor may be between the first molten carbonate fuel cell and a second molten carbonate fuel cell. The method for optimizing emission reduction of a power generation system may further comprise passing the recovered H₂ stream unadsorbed by the first swing adsorption reactor to a second anode of the second molten carbonate fuel cell and/or passing the CO₂ adsorbed by the first swing adsorption reactor to a second cathode of the second molten carbonate fuel cell.

The method of this embodiment may additionally include passing the recovered H₂ stream unadsorbed by the first swing adsorption to the second anode of the second molten carbonate fuel cell, using the CO₂ adsorbed by the first swing adsorption reactor for producing carbonate ions in the second cathode of the second molten carbonate fuel cell, transporting the produced carbonate ions from the second cathode to the second anode through a second electrolyte of the second molten carbonate fuel cell, and generating power and an exhaust gas by reacting the recovered H₂ with the transported carbonate ions in the second anode of the second molten carbonate fuel cell.

The method may further include passing the exhaust gas of the of the second molten carbonate fuel cell to a second swing adsorption reactor, wherein the exhaust gas includes CO₂ at a temperature greater than 600° C., and adsorbing the CO₂ from the exhaust gas on a second adsorbent material of the second swing adsorption reactor. The step of passing the exhaust gas of the first molten carbonate fuel cell to the first swing adsorption reactor may include using an ejector to drive a motive fluid under a pressure that is higher than a pressure of the exhaust gas of the first molten carbonate fuel cell. The ejector may be placed at a far side of the first swing adsorption reactor with respect to the first molten carbonate fuel cell, and the ejector produces a vacuum inside the first swing adsorption reactor by driving the motive fluid.

In another embodiment of the present invention, a method for optimizing emission reduction of a power generation system, comprises: passing hydrocarbons to a reformer for processing and producing H₂, passing the produced H₂ to a first anode of a first molten carbonate fuel cell, passing CO₂ and O₂ to a first cathode of the first molten carbonate fuel cell for processing and producing carbonate ions, transporting the produced carbonate ions from the first cathode to the first anode through a first electrolyte of the first molten carbonate fuel cell, generating power and an exhaust gas by reacting the H₂ with the transported carbonate ions in the first anode of the first molten carbonate fuel cell, the exhaust gas including CO₂ at a temperature greater than 600° C., passing the exhaust gas of the of the first molten carbonate fuel cell to a first swing adsorption reactor, and adsorbing the CO₂ from the exhaust gas on a first adsorbent material of the first swing adsorption reactor, wherein the first swing adsorption reactor is between the first molten carbonate fuel cell and a second molten carbonate fuel cell.

BRIEF DESCRIPTION OF THE FIGURES

FIG. 1 schematically shows an example of a combined cycle system for generating electricity based on combustion of a carbon-based fuel.

FIG. 2 schematically shows a configuration for a pressure swing adsorption process.

FIG. 3 shows a graph of CO₂ purity versus steam purge content for Example 1.

FIG. 4 pictorially shows the cycle setup of a 3-vessel sorptive arrangement without interconnection between vessels.

FIG. 5 pictorially shows the cycle setup of a 4-vessel sorptive arrangement without interconnection between vessels.

FIG. 6 pictorially shows the cycle setup of a 4-vessel sorptive arrangement with some level of interconnection between vessels.

FIG. 7 schematically shows a conventional example of a combined cycle system for generating electricity by molten carbonate fuel cell technology while reducing emissions.

FIG. 8 schematically shows an example of a combined cycle system of the present invention that generates electricity by molten carbonate fuel cell technology while reducing emissions.

FIGS. 9(a) and (b) schematically show an example of using an ejector to facilitate feeding of fuel cell exhaust gas into a PSA reactor.

FIG. 10 schematically shows an example of using a HotPSA to adsorb a high temperature ULET exhaust gas.

DETAILED DESCRIPTION OF THE EMBODIMENTS

In various aspects, systems and methods are provided for power generation using molten carbonate fuel cells while reducing and/or mitigating emissions during power generation. In a combined cycle generation system, an emission gas from a fuel cell reaction from one molten carbonate fuel cell can be either be cycled back into the system or used for power generation of another fuel cell. The emission gas can thus be recycled for use as part of the input to power generation system. A portion of the recycled emission gas can be effectively separated to generate a high purity carbon dioxide (CO₂) stream, while reducing and/or minimizing the energy required for the separation, and without needing to reduce the temperature of the emission gas. This can allow another (e.g., the remaining) portion of the recycled emission gas, which can typically include hydrogen (H₂), to be used to generate additional electricity, e.g., without having to adjust the pressure and/or temperature of the emission gas to accommodate the conditions required for the CO₂ separation process. Thus, improved energy recovery can be realized from the combined cycle system, while also generating relatively high purity streams of CO₂ and H₂.

It can be desirable to perform the separation of CO₂ and H₂ while reducing/minimizing the amount of energy lost due to temperature/pressure reductions. For example, typical solvent methods for separating CO₂ and H₂ require a reduction in the temperature of the recycled emission gas. For such conventional solvent methods, heat exchangers can be used to transfer heat from the recycled emission gas and lower its temperature prior to separation.

In some aspects, the separation of CO₂ and H₂ can be performed by using a pressure swing adsorption (PSA) process to separate the CO₂ and H₂ at the temperature and pressure of the gas flow at an outlet of a molten carbonate fuel cell. Using pressure swing adsorption to perform the separation can allow for recovery, for example, of at least about 60% (such as at least about 65%, at least about 70%, at least about 75%, at least about 80%, at least about 85%, at least about 90%, at least about 95%, at least about 97%, at least about 98%, at least about 99%, at least about 99.3%, or at least about 99.5%) of the CO₂ in the recycled emission gas, for example.

A high temperature PSA can be performed using high temperature adsorbent materials capable of performing adsorption at high temperatures of emission gases departing the molten carbonate fuel cell (>600° C.). Moreover, a high temperature PSA can be performed using a process cycle avoiding the need for high temperature steam and/or another energy intensive purge gas. This can allow high purity CO₂ to be recovered while reducing/minimizing the amount of energy lost to capture of the emission gases.

Combined Cycle Process for Power Generation with Low Emissions

In various aspects, systems and methods are provided for generating power while controlling and/or capturing the emissions produced during power generation. One goal of power generation is to use input feeds (such as fuels) as efficiently as possible, so that power generation can be increased/optimized for a given amount of fuel and/or of equipment. Based on the conditions for effective power generation, a goal for control/capture of emissions can be to provide effective capture of emissions while reducing/minimizing the changes to the conditions for power generation.

As used herein, the term “stoichiometric combustion” refers to a combustion reaction having a volume of reactants comprising a fuel and an oxidizer and a volume of products formed by combusting the reactants, where substantially the entire volume of the reactants is used to form the products. As used herein, the term “substantially stoichiometric combustion” refers to a combustion reaction having a molar ratio of combustion fuel to oxygen ranging from plus to minus about 10%, e.g., from about plus to minus about 5%, of the oxygen required for a stoichiometric ratio. For example, the stoichiometric ratio of fuel to oxygen for methane is 1:2 (CH₄+2O₂⇄CO₂+2H₂O), whereas propane should have a stoichiometric ratio of fuel to oxygen of 1:5. Another way of measuring substantially stoichiometric combustion can be as a ratio of oxygen supplied to oxygen required for stoichiometric combustion, e.g., from about 0.9:1 to about 1.1:1 or from about 0.95:1 to about 1.05:1.

In some aspects, the processes described herein can be used to produce ultra low emission electric power and CO₂ for EOR, enhanced hydrocarbon recovery (EHR), and/or sequestration/capture applications; in such cases, the process conditions for EOR/EHR may be similar to sequestration/capture application or may be slightly different. In one or more aspects, a mixture of oxygen-enriched gas (e.g., enriched air) and fuel can be stoichiometrically or substantially stoichiometrically combusted and simultaneously mixed with a stream of recycled exhaust gas. The stream of recycled exhaust gas, generally including products of combustion such as CO₂, can be used as a diluent to control, adjust, and/or otherwise moderate the temperature of combustion and the exhaust that enters the succeeding expander. As a result of using oxygen enrichment, the recycled exhaust gas can have an increased CO₂ content, thereby allowing the expander to operate at even higher expansion ratios for the same inlet and discharge temperatures, thereby producing significantly increased power.

Combustion in commercial gas turbines at stoichiometric conditions or substantially stoichiometric conditions (e.g., “slightly rich” combustion) can prove advantageous to eliminate the cost of excess oxygen removal. By cooling the exhaust and condensing the water out of the cooled exhaust stream, a relatively high content CO₂ exhaust stream can be produced. While a portion of the recycled exhaust gas can be utilized for temperature moderation in a closed Brayton cycle, a remaining purge stream can be used for EOR and/or enhanced hydrocarbon recovery applications and/or electric power can be produced with little or no sulfur oxides (SO_(x)), nitrogen oxides (NO_(x)), and/or CO₂ being emitted to the atmosphere. The result of this process can include the production of power in three separate cycles and the manufacturing of additional CO₂. In some aspects, performing stoichiometric combustion can allow for generation of an exhaust stream consisting substantially of CO₂, H₂O, and N₂. An exhaust stream consisting substantially of CO₂, H₂O, and N₂ is defined as an exhaust stream that contains about 5 mol % or less of other gas molecules, e.g., about 2.5 mol % or less or about 1 mol % or less.

FIG. 1 depicts a schematic of an illustrative integrated system 100 for power generation and CO₂ recovery using a combined-cycle arrangement, according to one or more embodiments. In at least one embodiment, the power generation system 100 can include a gas turbine system 102 characterized as a power-producing, closed Brayton cycle. The gas turbine system 102 can have a first or main compressor 104 coupled to an expander 106 via a shaft 108. The shaft 108 can be any mechanical, electrical, and/or other power coupling, thereby allowing a portion of the mechanical energy generated by the expander 106 to drive the main compressor 104. In at least one embodiment, the gas turbine system 102 can be a standard gas turbine, where the main compressor 104 and expander 106 form the compressor and expander ends, respectively. In other embodiments, however, the main compressor 104 and expander 106 can be individualized components in the system 102.

The gas turbine system 102 can also include a combustion chamber 110 configured to combust a fuel introduced via line 112 mixed with an oxidant introduced via line 114. In one or more embodiments, the fuel in line 112 can include any suitable hydrocarbon gas or liquid, such as natural gas, methane, ethane, naphtha, butane, propane, syngas, diesel, kerosene, aviation fuel, coal derived fuel, bio-fuel, oxygenated hydrocarbon feedstock, or any combinations thereof. The oxidant via line 114 can be derived from a second or inlet compressor 118 fluidly coupled to the combustion chamber 110 and adapted to compress a feed oxidant introduced via line 120. In one or more embodiments, the feed oxidant in line 120 can include atmospheric air, enriched air, or combinations thereof. When the oxidant in line 114 includes a mixture of atmospheric air and enriched air, the enriched air can be compressed by the inlet compressor 118 before and/or after being mixed with the atmospheric air. The enriched air can have an overall oxygen concentration of at least about 30 vol %, e.g., at least about 35 vol %, at least about 40 vol %, at least about 45 vol %, at least about 50 vol %, from about 30 vol % to about 70 vol %, from about 30 vol % to about 65 vol %, from about 30 vol % to about 60 vol %, from about 30 vol % to about 55 vol %, from about 30 vol % to about 50 vol %, from about 35 vol % to about 70 vol %, from about 35 vol % to about 65 vol %, from about 35 vol % to about 60 vol %, from about 35 vol % to about 55 vol %, from about 35 vol % to about 50 vol %, from about 40 vol % to about 70 vol %, from about 40 vol % to about 65 vol %, from about 40 vol % to about 60 vol %, from about 40 vol % to about 55 vol %, from about 40 vol % to about 50 vol %, from about 45 vol % to about 70 vol %, from about 45 vol % to about 65 vol %, from about 45 vol % to about 60 vol %, from about 45 vol % to about 55 vol %, from about 45 vol % to about 50 vol %, from about 50 vol % to about 70 vol %, from about 50 vol % to about 65 vol %, or from about 50 vol % to about 60 vol %.

The enriched air can be derived from any one or more of several sources, including implementing various technologies upstream of the inlet compressor 118 to produce the enriched air. For example, the enriched air can be derived from such separation technologies as membrane separation, pressure swing adsorption, temperature swing adsorption, nitrogen plant-byproduct streams, and/or combinations thereof. The enriched air can additionally or alternately be derived from an air separation unit (ASU), such as a cryogenic ASU, for producing nitrogen for pressure maintenance or other purposes. The reject stream from the ASU can be rich in oxygen, e.g., having an overall oxygen content from about 50 vol % to about 70 vol %. This reject stream can be used as at least a portion of the enriched air and subsequently diluted, if needed, with unprocessed atmospheric air to obtain the desired oxygen concentration for the application.

As will be described in more detail below, the combustion chamber 110 can also receive a compressed recycle exhaust gas in line 144, including an exhaust gas recirculation primarily having CO₂ and nitrogen components. The compressed recycle exhaust gas in line 144 can be derived from the main compressor 104 and adapted to help facilitate a stoichiometric or substantially stoichiometric combustion of the compressed oxidant in line 114 and fuel in line 112 by moderating the temperature of the combustion products. As can be appreciated, recirculating the exhaust gas can serve to increase the CO₂ concentration in the exhaust gas.

An exhaust gas in line 116 directed to the inlet of the expander 106 can be generated as a product of combustion of the fuel in line 112 and the compressed oxidant in line 114, in the presence of the compressed recycle exhaust gas in line 144. In at least one embodiment, the fuel in line 112 can be primarily natural gas, thereby generating a discharge or exhaust gas via line 116 that can include volumetric portions of vaporized water, CO₂, nitrogen, nitrogen oxides (NO_(x)), and sulfur oxides (SO_(x)). In some embodiments, a small portion of unburned fuel in line 112 or other compounds can also be present in the exhaust gas in line 116 due to combustion equilibrium limitations. As the exhaust gas in line 116 expands through the expander 106, it can generate mechanical power to drive the main compressor 104, an electrical generator, and/or other facilities, and can also produce a gaseous exhaust in line 122 having a heightened CO₂ content resulting from the influx of the compressed recycle exhaust gas in line 144. In some implementations, the expander 106 may be adapted to produce additional mechanical power that may be used for other purposes.

Additionally or alternately, the power generation system 100 can include an exhaust gas recirculation (EGR) system 124, which can include a heat recovery steam generator (HRSG) 126, or similar device, fluidly coupled to a steam gas turbine 128. In at least one embodiment, the combination of the HRSG 126 and the steam gas turbine 128 can be characterized as a power-producing closed Rankine cycle. In combination with the gas turbine system 102, the HRSG 126 and the steam gas turbine 128 can form part of a combined-cycle power generating plant, such as a natural gas combined-cycle (NGCC) plant. The gaseous exhaust in line 122 can be introduced to the HRSG 126 in order to generate steam via line 130 and a cooled exhaust gas in line 132. Additionally or alternately, the steam in line 130 can be sent to the steam gas turbine 128 to generate additional electrical power.

The cooled exhaust gas in line 132 can be sent to a first cooling unit 134 adapted to cool the cooled exhaust gas in line 132 and generate a cooled recycle gas stream 140. The first cooling unit 134 can include, for example, one or more contact coolers, trim coolers, evaporative cooling unit, or any combination thereof. The first cooling unit 134 can additionally or alternately be adapted to remove a portion of any condensed water from the cooled exhaust gas in line 132 via a water dropout stream 138. In at least one embodiment, the water dropout stream 138 may be routed to the HRSG 126 via line 141 to provide a water source for the generation of additional steam in line 130 therein. Additionally or alternately, the water recovered via the water dropout stream 138 can be used for other downstream applications, such as supplementary heat exchanging processes.

In most embodiments, the cooled recycle gas stream 140 can be directed to a boost compressor 142. Cooling the cooled exhaust gas in line 132 in the first cooling unit 134 can reduce the power required to compress the cooled recycle gas stream 140 in the boost compressor 142. As opposed to a conventional fan or blower system, the boost compressor 142 can be configured to compress, and thus increase, the overall density of the cooled recycle gas stream 140, thereby directing a pressurized recycle gas in line 145 downstream, where the pressurized recycle gas in line 145 can thus exhibit an increased mass flow rate for the same volumetric flow. This can prove advantageous, since the main compressor 104 can be volume-flow limited, and directing more mass flow through the main compressor 104 can result in higher discharge pressures, thereby translating into higher pressure ratios across the expander 106. Higher pressure ratios generated across the expander 106 can allow for higher inlet temperatures and, therefore, an increase in expander 106 power and/or efficiency. As can be appreciated, this may prove advantageous, since the CO₂-rich exhaust gas in line 116 can generally maintain a higher specific heat capacity.

Since the suction pressure of the main compressor 104 can typically be a function of its suction temperature, a cooler suction temperature can generally demand less power to operate the main compressor 104 for the same mass flow. Consequently, the pressurized recycle gas in line 145 can optionally be directed to a second cooling unit 136, e.g., which can include one or more direct contact coolers, trim coolers, evaporative cooling units, or any combination thereof. In at least one embodiment, the second cooling unit 136 can serve as an aftercooler adapted to remove at least a portion of the heat of compression generated by the boost compressor 142 on the pressurized recycle gas in line 145. The second cooling unit 136 can additionally or alternately extract additional condensed water via a water dropout stream 143. In some such embodiments, the water dropout streams 138,143 can converge into stream 141 and may or may not be routed to the HRSG 126 to generate additional steam via line 130 therein. While only first and second cooling units 134,136 are depicted herein, any desired number of cooling units can be employed to suit a variety of applications, without departing from the scope of the disclosure.

The main compressor 104 can be configured to receive and compress the pressurized recycle gas in line 145 to a pressure nominally at or above the pressure of the combustion chamber 110, thereby generating the compressed recycle exhaust gas in line 144. As can be appreciated, cooling the pressurized recycle gas in line 145 in the second cooling unit 136 after compression in the boost compressor 142 can allow for an increased volumetric mass flow of exhaust gas into the main compressor 104. Consequently, this can reduce the amount of power required to compress the pressurized recycle gas in line 145 to a predetermined pressure.

In many embodiments, a purge stream 146 can be recovered from the compressed recycle exhaust gas in line 144 and subsequently treated in a CO₂ separator 148 to capture CO₂ at an elevated pressure via line 150. Preferably, the CO₂ separator can be a pressure swing adsorption unit, as described in further detail below. The separated CO₂ in line 150 can be used for sales, used in another processes requiring CO₂, and/or further compressed and injected into a terrestrial reservoir for EOR, EHR, sequestration, or another purpose. Because of the stoichiometric or substantially stoichiometric combustion of the fuel in line 112 combined with a boosted pressure from the boost compressor 142, the CO₂ partial pressure in the purge stream 146 can be much higher than in conventional gas turbine exhausts.

A residual stream 151, essentially depleted of CO₂ and consisting primarily of nitrogen, can additionally or alternately be derived from the CO₂ separator 148. In some embodiments, the residual stream 151 can be introduced to a gas expander 152 to provide power and an expanded depressurized gas, or exhaust gas, via line 156. The expander 152 can be, for example, a power-producing nitrogen expander. As depicted, the gas expander 152 can be optionally coupled to the inlet compressor 118 through a common shaft 154 or other mechanical, electrical, or other power coupling, thereby allowing a portion of the power generated by the gas expander 152 to drive the inlet compressor 118. However, during startup of the system 100 and/or during normal operation, when the gas expander 152 is unable to supply all the required power to operate the inlet compressor 118, at least one motor 158, such as an electric motor, can be used synergistically with the gas expander 152. For instance, the motor 158 can be sensibly sized such that, during normal operation of the system 100, the motor 158 can be configured to supply the power short-fall from the gas expander 152. In other embodiments, however, the gas expander 152 can be used to provide power to other applications, and not directly coupled to the inlet compressor 118. For example, there may be a substantial mismatch between the power generated by the expander 152 and the requirements of the compressor 118. In such cases, the expander 152 could be adapted to drive a smaller (or larger) compressor (not shown) that may demand less (or more) power.

An expanded depressurized gas in line 156, primarily consisting of dry nitrogen gas, can be discharged from the gas expander 152. In at least one embodiment, the combination of the gas expander 152, inlet compressor 118, and CO₂ separator 148 can be characterized as an open Brayton cycle, and/or a third power-producing component of the power generation system 100. Conventional systems and methods of expanding the nitrogen gas in the residual stream 151, and variations thereof, are believed to be known in the art and are thus not discussed herein.

Additionally or alternately, gas expander 152 can be replaced and/or complemented with a downstream compressor 158. At least a portion (and up to all) of the residual stream 151 can be compressed in a downstream compressor to generate a compressed exhaust gas via line 160, which can be suitable for injection into a reservoir for pressure maintenance applications. In applications where methane gas is typically reinjected into hydrocarbon wells to maintain well pressures, compressing the residual stream 151 may prove advantageous. For example, the pressurized nitrogen gas in line 160 can instead be injected into the hydrocarbon wells, and any residual methane gas can be sold or otherwise used as fuel in related applications, such as in line 112.

By using enriched air as the compressed oxidant in line 114 and pressurizing the exhaust gas in the boost compressor 142, the power generation system 100 can achieve higher concentrations of CO₂ in the exhaust gas, thereby allowing for more effective CO₂ separation and capture. Embodiments disclosed herein, for example, can effectively increase the concentration of CO₂ in the exhaust gas in line 116 to CO₂ concentrations ranging from about 10 vol % to about 20 vol %. To achieve such CO₂ concentrations, the combustion chamber 110 can be adapted to stoichiometrically or substantially stoichiometrically combust an incoming mixture of fuel in line 112 and compressed oxidant in line 114, where the compressed oxidant in line 114 can include a stream having an oxygen content greater than about 21 vol %, e.g., enriched air, such as having an overall oxygen concentration of about 30 vol %, about 35 vol %, about 40 vol %, about 45 vol %, or about 50 vol %.

In order to moderate the temperature of the stoichiometric combustion and meet expander 106 inlet temperature and component cooling requirements, a portion of the exhaust gas with increased CO₂ content derived from the compressed recycle exhaust gas in line 144 can be injected into the combustion chamber 110 as a diluent. Thus, embodiments of the disclosure can essentially eliminate excess oxygen from the exhaust gas in line 116, while advantageously increasing its CO₂ concentration, e.g., up to about 20 vol % or optionally higher. As such, the gaseous exhaust in line 122 can have less than about 3.0 vol % oxygen, for example less than about 1.0 vol % oxygen, less than about 0.1 vol % oxygen, or less than about 0.01 vol % oxygen.

At least one benefit of having an increased CO₂ concentration can be that the expander 106 can be operated at an even higher expansion ratio for the same inlet and discharge temperatures, and can thereby produce increased power. This can be due to the higher heat capacity of CO₂ relative to nitrogen found in ambient air. In one or more aspects, the expansion ratio of the expander 106 can be increased from about 17.0 to about 20.0, corresponding to about 10 vol % and about 20 vol % CO₂ recycle streams, respectively. For example, enriched air having about 35 vol % oxygen can be used in order to achieve the about 20 vol % in the CO₂ recycle stream.

Additional or alternate benefits of having an increased CO₂ concentration in the recycle gas can include, but are not limited to, an increased concentration of CO₂ in the extracted purge stream 146 used for CO₂ separation. Because of its increased CO₂ concentration, the purge stream 146 need not be as large in order to extract the required amounts of CO₂. For example, the equipment handling extraction for CO₂ separation can be smaller, including its piping, heat exchangers, valves, absorber towers, etc. Moreover, increased concentrations of CO₂ can improve the performance of CO₂ removal technology, including using low-energy separation processes, such as employing less energy-intensive solvents that would otherwise be untenable. Consequently, capital expenditures for capturing CO₂ can be dramatically lowered.

An example of operation of the system 100 will now be discussed. As should be appreciated, specific temperatures and pressures achieved/experienced in the various components of any of the embodiments disclosed herein can change depending on, among other factors, the purity of the oxidant used and/or the specific makes and/or models of expanders, compressors, coolers, etc. Accordingly, it should be appreciated that the particular data described herein is for illustrative purposes only and should not be construed as the only interpretation thereof. In an embodiment, the inlet compressor 118 can provide compressed oxidant in line 114 at a pressure between about 280 psia (about 1.9 MPa) and about 300 psia (about 2.1 MPa). Also contemplated herein, however, is aeroderivative gas turbine technology, which can produce and consume pressures of up to about 750 psia (about 5.2 MPa) and higher.

The main compressor 104 can be configured to recycle and compress recycled exhaust gas into the compressed recycle exhaust gas in line 144 at a pressure nominally at or above the combustion chamber 110 pressure, and to use a portion of that recycled exhaust gas as a diluent in the combustion chamber 110. Because amounts of diluent needed in the combustion chamber 110 can depend on the purity of the oxidant used for stoichiometric combustion or the particular model/design of expander 106, a ring of thermocouples and/or oxygen sensors (not shown) can be disposed on the outlet of the expander 106. In operation, the thermocouples and/or sensors can be adapted to regulate and determine the volume of exhaust gas required as diluent needed to cool the products of combustion to the required expander inlet temperature, and also to provide feedback to regulate the amount of oxidant being injected into the combustion chamber 110. Thus, in response to the heat requirements detected by the thermocouples and/or the oxygen levels detected by the oxygen sensors, the volumetric mass flow of compressed recycle exhaust gas in line 144 and compressed oxidant in line 114 can be manipulated up or down to track the demand.

In at least one embodiment, a pressure drop of about 12-13 psi (about 83-90 kPa) can be experienced across the combustion chamber 110 during stoichiometric or substantially stoichiometric combustion. Combustion of the fuel in line 112 and the compressed oxidant in line 114 can generate temperatures between about 2000° F. (about 1093° C.) and about 3000° F. (about 1649° C.) and pressures ranging from about 250 psia (about 1.7 MPa) to about 300 psia (about 2.1 MPa). As described above, because of the increased mass flow and higher specific heat capacity of the CO₂-rich exhaust gas derived from the compressed recycle exhaust gas in line 144, higher pressure ratios can be achieved across the expander 106, thereby allowing for higher inlet temperatures and increased expander 106 power.

The gaseous exhaust in line 122 exiting the expander 106 can exhibit pressures at or near ambient, e.g., about 13-17 psia (about 90-120 kPa). The temperature of the gaseous exhaust in line 122 can be from about 1225° F. (about 663° C.) to about 1275° F. (about 691° C.) before passing through the HRSG 126 to generate steam in line 130 and a cooled exhaust gas in line 132. In one or more embodiments, the cooling unit 134 can reduce the temperature of the cooled exhaust gas in line 132, thereby generating the cooled recycle gas stream 140 having a temperature between about 32° F. (about 0° C.) and about 120° F. (about 49° C.). As can be appreciated, such temperatures can fluctuate, e.g., depending on wet bulb temperatures during specific seasons in specific locations around the globe.

According to one or more embodiments, the boost compressor 142 can be configured to elevate the pressure of the cooled recycle gas stream 140 to a pressure ranging from about 17 psia (about 120 kPa) to about 21 psia (about 140 kPa). As a result, the main compressor 104 can eventually receive and compress a recycled exhaust with a higher density and increased mass flow, thereby allowing for a substantially higher discharge pressure while maintaining the same or similar pressure ratio. In order to further increase the density and mass flow of the recycle exhaust gas, the pressurized recycle gas in line 145 discharged from the boost compressor 142 can then be further cooled in the optional second cooling unit 136, which can, in some embodiments, be configured to reduce the pressurized recycle gas temperature in line 145 to about 105° F. (about 41° C.) before being directed to the main compressor 104.

Additionally or alternately, the temperature of the compressed recycle exhaust gas in line 144 discharged from the main compressor 104, and consequently the temperature of the purge stream 146, can be about 800° F. (about 427° C.), with a pressure of around 280 psia (about 1.9 MPa). The addition of the boost compressor 142 and the stoichiometric combustion of enriched air can increase the CO₂ purge pressure in the purge stream 146, which can lead to improved solvent treating performance in the CO₂ separator 148 due to the higher CO₂ partial pressure.

Swing Adsorber Processes—Overview

Pressure swing adsorption (PSA) relies on swinging or cycling pressure over a bed of adsorbent through a range of values. In PSA processes, a gaseous mixture is conducted under pressure for a period of time over a first bed of a solid sorbent that is selective, or relatively selective, for one or more components, usually regarded as a contaminant, to be removed from the gaseous mixture. For example, a feed can be introduced into a PSA apparatus at a feed pressure. At the feed pressure, one or more of the components (gases) in the feed can be selectively (or relatively selectively) (ad)sorbed, while one or more other components (gases) can pass through with lower or minimal adsorption. A component (gas) that is selectively (ad)sorbed can be referred to as a “heavy” component of a feed, while a gas that is not selectively (ad)sorbed can be referred to as a “light” component of a feed. For convenience, a reference to the “heavy” component of the feed can refer to all components (gases) that are selectively (ad)sorbed, unless otherwise specified. Similarly, a reference to the “light” component can refer to all components (gases) that are not selectively (ad)sorbed, unless otherwise specified. After a period of time, the feed flow into the PSA apparatus can be stopped. The feed flow can be stopped based on a predetermined schedule, based on detection of breakthrough of one or more heavy components, based on (ad)sorption of the heavy component(s) corresponding to at least a threshold percentage of the total capacity of the (ad)sorbent, or based on any other convenient criteria. The pressure in the reactor can then be reduced to a desorption pressure that can allow the selectively (ad)sorbed component(s) (gas(es)) to be released from the (ad)sorbent. Optionally, one or more purge gases can be used prior to, during, and/or after the reduction in pressure to facilitate release of the selectively (ad)sorbed component(s) (gas(es)). Depending on its nature, a full PSA cycle can optionally be performed at a roughly constant temperature. As PSA is usually enabled by at least adsorption and usually occurs on gaseous components, the terms “adsorption”/“adsorbent” and “gas(es)” are used as descriptors in the instant specification and claims, without intending to be limiting in scope, even though “absorption”/“absorbent”/“sorbent”/“sorption” and “component(s)” may be more generally applicable.

Multiple beds can be used to enable a complete cycle, where typically every bed sequentially goes through the same cycle. When a first PSA reactor satisfies a condition, such as the adsorbent in the reactor becoming sufficiently saturated, the feed flow can be switched to a second reactor. The first PSA reactor can then be regenerated by having the adsorbed gases released. To allow for a continuous feed flow, a sufficient number of PSA reactors and/or adsorbent beds can be used so that the first PSA reactor is finished regenerating prior to at least one other PSA reactor satisfying the condition for switching reactors.

Swing Adsorber Processes—Process Cycle

In various aspects, a PSA reactor can be used for performing a separation on a stream containing CO₂ and N₂. An example of such a stream can include the exhaust stream from a combustion reaction for providing power for a gas turbine. Preferably, the exhaust stream can be the exhaust from a combustion reaction performed with a substantially stoichiometric composition with regard to the amount of oxygen and fuel. Prior to use as a feed for separation by PSA, the exhaust stream can undergo further processing, such as condensation to remove water, combustion to remove excess fuel, adsorption for removal of NO_(x) species, and/or other types of processing to remove components different from CO₂ and N₂. In some aspects, the portion of the exhaust stream used as the feed for the PSA reactor can have a water content of less than about 1.0 vol %, such as less than about 0.5 vol %. Additionally or alternately, the portion of the exhaust stream used as the feed for the PSA reactor can have an O₂ content of less than about 3.0 vol %, such as less than about 1.0 vol % or less than about 0.5 vol %. Further additionally or alternately, the feed into the PSA reactor can be substantially composed of CO₂ and N₂, where components of the input gas feed different from CO₂ and N₂ are present in an amount of about 1.0 vol % or less each, such as less than about 0.5 vol % each. Still further additionally or alternately, in a feed substantially composed of CO₂ and N₂, the combined vol % of components other than CO₂ and N₂ can be about 2.0 vol % or less, such as about 1.0 vol % or less or 0.5 vol % or less.

To perform a separation, a portion of the recycled exhaust stream can be introduced into a PSA reactor, such as a purge stream from the exhaust stream recycle loop. The portion of the recycled exhaust stream can be withdrawn from the exhaust recycle system at a location after the temperature and pressure of the recycled exhaust stream have been modified (and/or after desired temperature and pressure have been obtained) for use as part of the input gas flow to the combustion reaction. At such point in the recycle system, the exhaust stream can have a temperature from about 300° C. to about 600° C., e.g., from about 300° C. to about 550° C., from about 300° C. to about 500° C., from about 300° C. to about 450° C., from about 300° C. to about 400° C., from about 350° C. to about 600° C., from about 350° C. to about 550° C., from about 350° C. to about 500° C., from about 350° C. to about 450° C., from about 400° C. to about 600° C., from about 400° C. to about 550° C., from about 400° C. to about 500° C., from about 425° C. to about 600° C., from about 425° C. to about 550° C., from about 425° C. to about 500° C., from about 425° C. to about 460° C., from about 450° C. to about 600° C., from about 450° C. to about 550° C., from about 500° C. to about 600° C. Additionally or alternately, the temperature can be at least about 325° C., e.g., at least about 350° C., at least about 400° C., at least about 425° C., or at least about 440° C. Further additionally or alternately, the temperature can be about 575° C. or less, about 550° C. or less, about 500° C. or less, about 475° C. or less, about 460° C. or less, or about 440° C. or less. Still further additionally or alternately, the pressure of the recycled exhaust stream can be at least about 10 bara (about 1.0 MPa), e.g., at least about 15 bara (about 1.5 MPa), at least about 20 bara (about 2.0 MPa), at least about 25 bara (about 2.5 MPa), or at least about 30 bara (about 3.0 MPa). Yet further additionally or alternately, the pressure can be about 60 bara (6.0 MPa) or less, e.g., about 50 bara (about 5.0 MPa) or less, about 40 bara (about 4.0 MPa) or less, about 35 bara (about 3.5 MPa) or less, about 30 bara (about 3.0 MPa) or less, about 25 bara (about 2.5 MPa) or less, or about 22.5 bara (about 2.25 MPa) or less. In some alternative aspects, other locations for withdrawing the portion of the recycled exhaust stream can be selected, so long as the withdrawn portion has similar temperature and pressure values.

In some aspects, the recycled exhaust stream can be introduced into the PSA reactor at a separation temperature and a separation pressure that can correspond to the temperature and pressure of the recycle stream. However, some variation in temperature and/or pressure may occur between withdrawal of the portion of the recycled exhaust stream from the recycle system and introduction of the recycled exhaust stream into the PSA reactor. For example, the separation temperature for the portion of the recycled exhaust stream introduced into the PSA reactor can differ from the temperature in the recycle system by about 20° C. or less, e.g., by about 15° C. or less or by about 10° C. or less. It is noted that, although having the same temperature for the separation temperature and the temperature in the recycle stream (i.e., a difference of approximately 0° C.) is included within the enumerated differences, in some rare embodiments the temperature differences can optionally exclude 0° C. Additionally or alternately, the separation pressure for the portion of the recycled exhaust stream introduced into the PSA reactor can differ from the pressure in the recycle system by about 5 bar (0.5 MPa) or less, e.g., about 2 bara (0.2 MPa) or less, about 1 bara (about 0.1 MPa) or less, or about 0.5 bar (about 50 kPa) or less.

When the exhaust stream is introduced into the PSA reactor, the N₂ in the exhaust stream corresponds to a “light” component while the CO₂ corresponds to a “heavy” component. Thus, the N₂ can primarily pass through the reactor while the CO₂ can be selectively adsorbed within the reactor. The adsorption of CO₂ from the feed can result in a product N₂ stream. The feed can be passed through the PSA reactor until one or more pre-defined criteria is satisfied for switching the feed to another PSA reactor or otherwise stopping the flow of feed gas. Any convenient pre-defined criteria can be used. For example, the feed can be passed through the reactor for a specified time period. Additionally or alternately, the feed can be passed into the reactor until a breakthrough amount of CO₂ is detected in the product N₂ stream. Further additionally or alternately, the feed can be passed into the reactor until the amount of CO₂ that has entered the reactor is approximately equal to a threshold value of the adsorbent capacity of the reactor. In such a situation, for example, the feed can be passed into the reactor until the amount of CO₂ that has entered the reactor is equal to at least about 75% of the adsorbent capacity of the adsorbent material in the reactor, such as at least about 80%, at least about 85%, or at least about 90%. A typical PSA cycle can involve introducing feed into the reactor for about 30 seconds to about 300 seconds, e.g., for about 60 seconds to about 120 seconds.

The product N₂ stream can have a purity of at least about 85 vol %, e.g., at least about 88 vol %, at least about 90 vol %, at least about 92 vol %, at least about 93 vol %, at least about 94 vol %, at least about 95 vol %, at least about 96 vol %, at least about 97 vol %, or at least about 98 vol %. After optional removal of water, the product N₂ stream can have a purity of at least about 90 vol %, e.g., at least about 95 vol %, at least about 97 vol %, at least about 98 vol %, or at least 99 vol %. The pressure of the N₂ stream can be at least about 90%, e.g., at least about 95%, of the pressure of the recycled exhaust used as the input to the PSA reactor. Additionally or alternately, the pressure of the N₂ stream can differ from the separation pressure by less than about 0.5 bar (about 50 kPa), e.g., less than about 0.3 bar (about 30 kPa) or less than about 0.1 bar (about 10 kPa). Further additionally or alternately, the separation temperature for the portion of the recycled exhaust stream introduced into the PSA reactor can differ from the temperature of the N₂ product stream by about 20° C. or less, e.g., by about 15° C. or less or by about 10° C. or less. It is noted that, although having the same temperature for the separation temperature and the temperature in the recycle stream (i.e., a difference of approximately 0° C.) is included within the enumerated differences, in some rare embodiments the temperature differences can optionally exclude 0° C. Additionally or alternately, the temperature of the N₂ product stream can be at least the temperature of the PSA reactor (the adsorption temperature) during the adsorption portion of the separation cycle.

After the feed is stopped, the pressure in the PSA reactor can be reduced, e.g., using one or more blow down processes. In a blow down process, one or both sides of a PSA reactor can be opened to allow pressure to release in the form of a blow down gas stream. The blow down gas stream can generally include a majority portion of N₂ and can also typically include some CO₂. The amount of adsorbed CO₂ released in the blow down process(es) can depend on the nature of the adsorbent. In some conventional PSA reactors, the blow down gas stream can be exhausted from the feed input side of the reactor. Alternatively, one or more blow down gas streams can be exhausted from the product side of the reactor. For example, one option can include having an initial blow down process exiting from the product side of the reactor followed by a second blow down process allowing a gas stream to exit from both sides of the reactor. The blow down process(es) can reduce the pressure in the reactor to a value from about 0.9 bara (about 90 kPa) to about 3.0 bara (about 0.3 MPa), e.g., from about 1.0 bara (about 0.1 MPa) to about 3.0 bara (about 0.3 MPa), from about 1.1 bara (about 110 kPa) to about 3.0 bara (about 0.3 MPa), from about 1.3 bara (about 130 kPa) to about 3.0 bara (about 0.3 MPa), from about 0.9 bara (about 90 kPa) to about 2.6 bara (about 260 kPa), from about 1.0 bara (about 0.1 MPa) to about 2.6 bara (about 260 kPa), from about 1.1 bara (about 110 kPa) to about 2.6 bara (about 260 kPa), from about 1.3 bara (about 130 kPa) to about 2.6 bara (about 260 kPa), from about 0.9 bara (about 90 kPa) to about 2.0 bara (about 0.2 MPa), from about 1.0 bara (about 0.1 MPa) to about 2.0 bara (about 0.2 MPa), from about 1.1 bara (about 110 kPa) to about 2.0 bara (about 0.2 MPa), or from about 1.3 bara (about 130 kPa) to about 2.0 bara (about 0.2 MPa). In many embodiments, maintaining a pressure above atmospheric pressure in the reactor can assist with the adsorbent retaining CO₂ until a subsequent purge step when desorption is desired. In some embodiments, the length of time for the blow down processes can be from about 30 seconds to about 120 seconds.

In some aspects, the use of multiple blow down steps can be desirable for creating blow down streams that are easier to subsequently process. For example, during adsorption of CO₂ from an exhaust gas, a profile can typically develop in the reactor, with a higher content of non-adsorbed CO₂ near the back (input) end of the reactor and a lower content of CO₂ near the exit (front) end of the reactor. Based on this profile, a partial blow down from only the exit (front) end of the reactor can be used to produce a blow down output with a low CO₂ content. This initial blow down step can result in the higher CO₂ content near the back (input) end of the reactor being distributed more evenly throughout the reactor. As this occurs, it can then be more efficient to allow blow down output streams to exit from both ends of the reactor until the desired lower pressure can be achieved.

Preferably, a buffer gas is not introduced into the reactor during the time between stopping the flow of exhaust gas and starting the blow down process step(s). It can additionally or alternately be preferred that the blow down process step(s) can be performed without introducing an additional gas into the reactor. Avoiding the use of buffer gases and/or additional gases in the blow down steps can be desirable, because introduction of such gases after the flow of exhaust gas is stopped can typically result in further loss of value into a low value stream. For example, the output flows generated by the blow down step(s) are, by definition, lower pressure output flows relative to the product N₂ flow generated during the separation of the exhaust gas. Thus, the N₂ in the blow down output has typically lost much of its value from a power generation standpoint. Prior to further use, the N₂ in the blow down output flow can likely need to be recompressed. Adding additional N₂ after the flow of exhaust gas is stopped can typically only increase the amount of N₂ in this low value output flow. The blow down output flow can preferably be relatively low in CO₂, as it can generally be desirable to retain as much CO₂ as possible until the start of the subsequent purge step(s). Any CO₂ that exits the PSA reactor as part of a blow down stream represents additional CO₂ in a stream other than the desired CO₂ product stream. This additional CO₂, which can typically be in low concentration, can then need to be separately handled if it is desired to achieve as high an amount of carbon capture and recovery as possible. Thus, adding additional CO₂ here is also not typically desirable. Finally, adding a third gas different from CO₂, N₂, or H₂O may not be desirable, as introduction of such a gas can likely result in another component requiring separation.

After the blow down process(es), one or more purge gas flows can be used to remove the adsorbed CO₂ from the reactor. One option can include using a steam purge at a pressure at <1 bara (0.1 MPa) to assist in desorbing the CO₂. The steam purge can result in a product CO₂ output stream that can also include H₂O and a lesser amount of N₂. In some embodiments, the steam purge can last for about 25 seconds to about 60 seconds. After removal of water, the product CO₂ stream can have a purity of at least about 60%, e.g., at least about 65%, at least about 70%, at least about 75%, at least about 80%, at least about 85 vol %, at least about 90 vol %, at least about 92 vol %, at least about 94 vol %, at least about 95 vol %, at least about 96%, at least about 97%, at least about 98%, at least about 99%, at least about 99.3%, or at least about 99.5%. Additionally or alternately, the amount of CO₂ recovered can correspond to at least about 80 vol %, e.g., at least about 85 vol %, at least about 90 vol %, at least about 92 vol %, at least about 94 vol %, at least about 95 vol %, at least about 96%, at least about 97%, at least about 98%, at least about 99%, at least about 99.3%, or at least about 99.5%.

The amount of steam used in the steam purge can correspond to about 1.0 moles of water or less per mole of CO₂ in the feed, e.g., about 0.9 moles of water or less per mole of CO₂ in the feed, about 0.75 moles of water or less per mole of CO₂ in the feed, about 0.6 moles of water or less per mole of CO₂ in the feed, about 0.5 moles of water or less per mole of CO₂ in the feed, or about 0.4 moles of water or less per mole of CO₂ in the feed. Using less steam in the purge can be beneficial, because the amount of steam used can typically correspond to the amount of energy used for the CO₂ recovery. It is noted that the feed can typically contain at least 50 mol %, and often at least 75 mol %, of N₂. Thus, a comparison of the number of moles of water per total moles of gas (including both N₂ and CO₂) in the feed could produce still lower values, e.g., about 0.5 moles of water or less per mole of gas in the feed, about 0.4 moles of water or less per mole of gas in the feed, about 0.3 moles of water or less per mole of gas in the feed, about 0.25 moles of water or less per mole of gas in the feed, about 0.2 moles of water or less per mole of gas in the feed, about 0.15 moles of water or less per mole of gas in the feed, or about 0.1 moles of water or less per mole of gas in the feed. In such embodiments where there is a non-zero amount of steam used in the steam purge, the molar ratio of steam can be at least about 0.05 moles of water or less per mole of CO₂ in the feed (e.g., at least about 0.1 moles of water or less per mole of CO₂ in the feed, at least about 0.2 moles of water or less per mole of CO₂ in the feed, or at least about 0.3 moles of water or less per mole of CO₂ in the feed) and/or can be at least about 0.01 moles of water or less per mole of gas in the feed (e.g., at least about 0.02 moles of water or less per mole of gas in the feed, at least about 0.05 moles of water or less per mole of gas in the feed, or at least about 0.1 moles of water or less per mole of gas in the feed).

After the steam purge, a second nitrogen purge can optionally be used to remove water and any remaining CO₂ from the reactor. If the second purge is not used, the N₂ output stream in the next cycle may have higher water content, but otherwise the additional water is not believed to significantly impact the separation process. When a second nitrogen purge is used, it can be performed for about 10 seconds to about 45 seconds. After the steam purge and the optional second nitrogen purge, the reactor can then be repressurized to start the next separation cycle. The repressurization can be performed using the input feed (the recycled exhaust gas). Alternatively but preferably, a second purge, such as a second nitrogen purge, may not be performed, as such a second purge can represent a flow of buffer gas between the end of a purge step and the start of the next cycle of introducing exhaust gas for separation of N₂ and CO₂. For the reasons noted above, additional buffer gas steps can typically be undesirable, as such buffer steps can tend to result in additional volume for low value gas streams.

In an alternative embodiment, no steam is intentionally used as a purge gas in the desorption step of the PSA cycle. In such an alternative embodiment, if a purge gas is used at all, it would not intentionally include water but, for instance, could include a nitrogen purge (e.g., which can last for about 10 seconds to about 60 seconds, for about 10 seconds to about 50 seconds, for about 10 seconds to about 45 seconds, for about 15 seconds to about 60 seconds, for about 15 seconds to about 50 seconds, for about 15 seconds to about 45 seconds, for about 20 seconds to about 60 seconds, for about 20 seconds to about 50 seconds, for about 20 seconds to about 45 seconds, for about 25 seconds to about 60 seconds, for about 25 seconds to about 50 seconds, or for about 25 seconds to about 45 seconds). It should be understood that, even in such embodiments where no steam is intentionally added, some water/steam may nevertheless be present as an impurity/contaminant in any purge stream that is used. In such alternative embodiments, the use of no intentionally added steam can allow the CO₂ from the PSA process to be introduced directly into CO₂ compressors, e.g., for well injection in enhanced oil recovery processes, without the need for intermediate equipment such as condensers or heat exchangers. Similar recoveries and/or purities of CO₂ can be achieved in this alternative embodiment, though obviously water need not be removed to attain those recoveries and/or purities in such an embodiment.

An additional/alternate way of characterizing the desire to reduce/minimize the use of buffer gases can be by characterizing the total gas input flows into the PSA reactor during a process cycle. In the process described above, a full cycle of the process can correspond to passing an initial gas flow into the reactor for separation, adsorbing CO₂ from the input gas flow, recovering an N₂ product stream, blowing down the pressure in the reactor, and purging the reactor with steam, optionally including a second purge stream. If only a steam purge is used, the primary input gas flows into the PSA reactor during a process cycle can include the input gas flow of recycled exhaust gas and the steam purge. Preferably, these primary input flows can correspond to at least about 90 vol % of all gas flows into the PSA reactor, e.g., at least about 95 vol %, at least about 98 vol %, or at least about 99 vol %. If a second purge of N₂ or another second purge gas is used, then the primary input gas flows into the PSA reactor during a process cycle can include the input gas flow of recycled exhaust gas, the steam purge, and the second purge. Preferably, in such an embodiment, these primary input flows can correspond to at least about 90 vol % of all gas flows into the PSA reactor, e.g., at least about 95 vol %, at least about 98 vol %, or at least about 99 vol %.

FIG. 2 schematically shows an example of the gas flows into and out of a pressure swing adsorption unit suitable for use in separating an exhaust gas containing CO₂ and N₂. In FIG. 2, a pressure swing adsorption reactor 210 can receive a feed 205 containing CO₂ and N₂ for separation from a first or back side of the reactor. An N₂ product stream 202 can emerge from a second or forward side of the reactor. After a period of time, such as after breakthrough of the CO₂, the feed to the PSA can be stopped. One or more blow down processes can then be performed to reduce the pressure in the reactor. In FIG. 2, a forward blow down output 222 is shown. Optionally, either a back blow down or both a back and a forward blow down can be used in place of a forward blow down 222. A purge stream such as a steam purge 215 can then be introduced into the forward side of the reactor. The purge stream can assist in desorbing CO₂ from the adsorbent to produce the CO₂ product stream 212.

Though not specifically delineated in any of the Figures herein, the purified CO₂ output of the PSA process can be used in applications to recover oil, gas, or associated hydrocarbons from surface and/or subterranean deposits (e.g., conventional downhole oil/natural gas reservoirs, unconventional reservoirs/wells, shale oil deposits, shale gas deposits, tight gas sands, stranded gas deposits, fracking deposits, and the like). In many of these embodiments, relatively high CO₂ purities may be desirable, e.g., to increase efficiency and reduce costs associated with compressing impurities. Nevertheless, in certain of these embodiments, at least some part of enhanced oil/gas/hydrocarbon recovery can efficiently utilize lower CO₂ purities (e.g., from about 60% to about 85% or from about 60% to about 80%), and particularly can offer a distinct benefit from methods according to the invention that allow (increased) flexibility in CO₂ stream purity (e.g., a staged flood, wherein a sequence of different purities are used). Such flexible purity control can be aligned with a particular deposit/reservoir that can age over time and/or with the quality of a particular deposit.

Swing Adsorber Processes—Reactor Configuration

Beds of adsorbent can be arranged in any convenient manner to provide a flow path for gas, including axial and/or radial flow directions. The vessels holding the bed of adsorbent can be oriented with respect to flow direction in any convenient manner. One typical orientation for a PSA reactor can be to have a reactor where the direction of flow is aligned with the long or primary axis of the reactor. An example of this includes having PSA reactors in the form of vertical cylinders, where the input gas flow enters the cylinder through the top or bottom surface. While this can represent a conventional configuration for a PSA reactor relative to the direction of gas flow, scaling up a process to handle large flow volumes can pose difficulties. For example, a typical length to diameter ratio for a PSA reactor can be about 3 to 1, such as a reactor with a length of about 10 meters and a diameter of about 3 meters. Increasing the size of such a reactor can generally result in an increasing pressure drop across the reactor, which is typically not desirable. Thus, in order to scale up axial flow PSA reactors to handle large gas flows, multiple reactors can typically be used.

Additionally or alternately, a PSA reactor can be used where the long or primary axis of the reactor can be perpendicular to the direction of gas flow. For example, a PSA reactor can include parallel plates of adsorbent with the long axis of the parallel plates being perpendicular to the direction of gas flow. This can allow a much larger volume of adsorbent for capturing CO₂ to be placed in a single reactor with an axial flow orientation while reducing the distance the gas flow must travel to cross the reactor. As a result, this type of configuration can reduce the pressure drop across a PSA reactor while still handling large volumes of input gas flow. This can reduce the number of separate reactors required to handle a large flow.

As an example, a horizontally oriented PSA reactor can have parallel plates of adsorbent of approximately rectangular shape, e.g., with a long axis of about 30 meters and a short axis of about 4 meters. In the horizontal configuration, the input gas can be introduced to flow across the short axis. This can be accomplished by introducing the input gas at several locations along a side of the reactor corresponding to the long axis. Flow distributors can then be used inside the PSA reactor so that the input gas can be distributed along the entire length of the long axis. For large gas flow rates, the pressure drop from distributing the feed across the long axis of a PSA reactor can be lower than attempting to force such a large gas flow rate through a plurality of conventional axial flow PSA reactors.

Swing Adsorber Processes—Adsorbent Materials

In various aspects, a swing adsorption process can be performed to separate N₂ from CO₂ at a temperature and pressure beneficial for other aspects of the combined power generation process. For example, the recycled exhaust gas can have a temperature from about 300° C. to about 600° C. (e.g., from about 300° C. to about 550° C., from about 300° C. to about 500° C., from about 300° C. to about 450° C., from about 300° C. to about 400° C., from about 350° C. to about 600° C., from about 350° C. to about 550° C., from about 350° C. to about 500° C., from about 350° C. to about 450° C., from about 400° C. to about 600° C., from about 400° C. to about 550° C., from about 400° C. to about 500° C., from about 425° C. to about 600° C., from about 425° C. to about 550° C., from about 425° C. to about 500° C., from about 425° C. to about 460° C., from about 450° C. to about 600° C., from about 450° C. to about 550° C., or from about 500° C. to about 600° C.; additionally or alternately, the temperature can be at least about 325° C., e.g., at least about 350° C., at least about 400° C., at least about 425° C., or at least about 440° C.; further additionally or alternately, the temperature can be about 575° C. or less, about 550° C. or less, about 500° C. or less, about 475° C. or less, about 460° C. or less, or about 440° C. or less) and a pressure from about 10 bara (about 1.0 MPa) to about 60 bara (about 6.0 MPa) (e.g., a pressure of at least about 10 bara (about 1.0 MPa), at least about 15 bara (about 1.5 MPa), at least about 20 bara (about 2.0 MPa), at least about 25 bara (about 2.5 MPa), or at least about 30 bara (about 3.0 MPa) and/or a pressure of about 60 bara (6.0 MPa) or less, about 50 bara (about 5.0 MPa) or less, about 40 bara (about 4.0 MPa) or less, about 35 bara (about 3.5 MPa) or less, about 30 bara (about 3.0 MPa) or less, about 25 bara (about 2.5 MPa) or less, or about 22.5 bara (about 2.25 MPa) or less).

As noted above, one goal of the separation process can be to perform the separation under conditions compatible with the power generation process. Thus, it can be desirable to perform the separation at approximately the temperature and pressure of the recycled exhaust gas.

In order to accomplish a separation at the conditions of the recycled exhaust gas, the adsorbent material in the pressure swing adsorber reactor should generally be effective under such conditions.

One example of a suitable adsorbent includes a mixed metal oxide adsorbent, such as an adsorbent including a mixture of an alkali metal carbonate and an alkaline earth metal oxide and/or a transition metal oxide. Examples of suitable alkali metal carbonates can include, but are not limited to, a carbonate of lithium, sodium, potassium, rubidium, cesium, or a combination thereof, e.g., a carbonate of lithium, sodium, potassium, or a combination thereof. Examples of suitable alkaline earth metal oxides can include, but are not limited to, oxides of magnesium, calcium, strontium, barium, or a combination thereof, e.g., oxides of magnesium and/or calcium. Some examples of suitable transition metal oxides can include, but are not limited to, oxides of lanthanide series metals, such as lanthanum, and/or of transition metals that can form oxides with the metal in a +2 or +3 oxidation state (such as yttrium, iron, zinc, nickel, vanadium, zirconium, cobalt, or a combination thereof).

In some aspects, the carbonate can be selected independently from the oxide in the mixed metal oxide. In such aspects, the carbonate can include, consist essentially of, or be lithium carbonate, sodium carbonate, potassium carbonate, rubidium carbonate, and/or cesium carbonate (e.g., lithium carbonate, sodium carbonate, and/or potassium carbonate; lithium carbonate and/or potassium carbonate; lithium carbonate and/or sodium carbonate; or sodium carbonate and/or potassium carbonate).

In aspects where the carbonate is selected independently from the oxide, the oxide can be an alkaline earth oxide, a transition metal oxide, a combination of two or more alkaline earth oxides, a combination of two or more transition metal oxides, or a combination of oxides including at least one alkaline earth oxide and at least one transition metal oxide. In aspects where the independently selected oxide includes one or more alkaline earth oxides, a suitable alkaline earth oxide can include, consist essentially of, or be magnesium oxide, calcium oxide, strontium oxide, and/or barium oxide, e.g., including at least magnesium oxide and/or calcium oxide.

In aspects where the independently selected oxide includes one or more transition metal oxides, suitable transition metals can include, consist essentially of, or be one or more transition metals that can form oxides with the metal in a +2 or +3 oxidation state (e.g., yttrium oxide, iron oxide, zinc oxide, nickel oxide, vanadium oxide, cobalt oxide, zirconium oxide, lanthanum oxide, other oxides of lanthanide metals, and/or a combination thereof). One preferred option includes a transition metal oxide selected from lanthanum oxide and/or zirconium oxide. Another option includes a metal oxide selected from lanthanum oxide, yttrium oxide, zirconium oxide, and/or zinc oxide. Yet another option includes a metal oxide selected from nickel oxide, cobalt oxide, and/or iron oxide. Mixtures within each of these options and/or across options are also contemplated, such as mixtures of lanthanum oxide with zinc oxide and/or vanadium oxide; mixtures of lanthanum oxide with iron oxide, cobalt oxide, and/or nickel oxide; mixtures of zirconium oxide with yttrium oxide, zinc oxide, and/or vanadium oxide; and mixtures of zirconium oxide with iron oxide, cobalt oxide, and/or nickel oxide.

In aspects where the independently selected oxide includes one or more alkali metal oxides and one or more transition metal oxides, suitable alkali metal oxides can include, consist essentially of, or be magnesium oxide, calcium oxide, strontium oxide, and/or barium oxide, while suitable transition metals can include, consist essentially of, or be transition metals that can form oxides with the metal in a +2 or +3 oxidation state, such as yttrium oxide, iron oxide, zinc oxide, nickel oxide, vanadium oxide, cobalt oxide, zirconium oxide, lanthanum oxide, and/or other lanthanide oxides. Each of these alkali metal oxides and transition metal oxides can be independently selected individually or in any combination of multiple transition metal oxides. Examples of mixtures can include, consist essentially of, or be a mixture of oxides where at least one oxide is lanthanum oxide, zirconium oxide, and/or magnesium oxide; a mixture of oxides where the mixture includes at least two of lanthanum oxide, zirconium oxide, and magnesium oxide; a mixture of oxides where one oxide is magnesium oxide and/or calcium oxide; and/or a mixture of oxides where at least one oxide is lanthanum oxide, yttrium oxide, and/or zirconium oxide.

In some alternative aspects, a mixed metal oxide can include an alkaline earth carbonate in combination with a transition metal oxide. In such aspects, the alkaline earth carbonate can include, consist essentially of, or be magnesium carbonate and/or calcium carbonate. Additionally or alternately, the alkaline earth carbonate can be present in a mixture with an alkali metal carbonate. Examples of such carbonate mixtures can include, consist essentially of, or be mixtures of lithium carbonate with magnesium carbonate, lithium carbonate with calcium carbonate, potassium carbonate with magnesium carbonate, potassium carbonate with calcium carbonate, sodium carbonate with magnesium carbonate, and sodium carbonate with calcium carbonate (e.g., lithium carbonate with magnesium carbonate or potassium carbonate with magnesium carbonate). In such aspects, suitable transition metals can include, consist essentially of, or be transition metals that can form oxides with the metal in a +2 or +3 oxidation state, such as yttrium oxide, iron oxide, zinc oxide, nickel oxide, vanadium oxide, cobalt oxide, zirconium oxide, lanthanum oxide, other lanthanide oxides, and/or a combination thereof. Each of these alkaline earth carbonates and transition metal oxides can be independently selected individually or in any combination of multiple alkaline earth carbonates and/or multiple transition metal oxides. For the transition metal oxide, one preferred option can include a transition metal oxide selected from lanthanum oxide or zirconium oxide. Another option can include a metal oxide selected from lanthanum oxide, yttrium oxide, zirconium oxide, and/or zinc oxide. Yet another option can include a metal oxide selected from nickel oxide, cobalt oxide, and/or iron oxide. Mixtures within each of these options and/or across options are also contemplated, such as mixtures of oxides where at least one oxide is lanthanum oxide and/or zirconium oxide; mixtures of lanthanum oxide with zinc oxide and/or vanadium oxide; mixtures of lanthanum oxide with iron oxide, cobalt oxide, and/or nickel oxide; mixtures of zirconium oxide with yttrium oxide, zinc oxide, and/or vanadium oxide; and/or mixtures of zirconium oxide with iron oxide, cobalt oxide, and/or nickel oxide.

Additional or alternative materials can include hydrotalcites.

Several examples of pressure swing adsorption are described below.

EXAMPLE 1 Performing PSA on a Portion of Recycled Exhaust

FIG. 3 shows how the recovery (represented as a molar ratio) and purity (represented as a percent) trade off against steam usage for exemplary PSA configurations that can further illustrate the invention. For each of the cases shown, the feed composition and pressure are similar to those described in the examples above. In FIG. 3, either ˜3-vessel or ˜4-vessel PSA configurations are illustrated for typical adsorbent properties of mass transfer and adsorbent capacity. As seen in FIG. 3, higher purity CO₂ (>95 vol %) can be achievable with increased steam usage, but still less than ˜1.0 mol/mol of CO₂ recovered. In addition, the use of ˜4 PSA vessels can allow a reduced pressure swing via the use of a higher relative purge pressure (>about 1.35-1.5 bara, typically from ˜2 bara to ˜3 bara), which can reduce both equipment and energy requirements for subsequent CO₂ compression prior to use in EOR facilities. Further improvement of the ˜4-vessel recoveries shown in FIG. 3 can be obtained using configurations where the vessels are in fluid communication with each other.

EXAMPLE 2 3-Vessel PSA Configuration using a Portion of Recycled Exhaust

FIG. 4 shows a 3-vessel configuration that is illustrative but not limiting of one possible arrangement of adsorber vessels within the scope of the invention. In this configuration, one vessel is on feed and hence making nitrogen product at all times. The other two vessels can be undergoing other steps of the cycle, e.g., that regenerate the adsorbent, whilst the first vessel is making product at relatively high pressure. In this particular 3-vessel arrangement, there is no fluid connection (i.e., exchange of gas streams) between the vessels, which reduces cycle complexity as well as equipment such as interconnecting piping and valving.

EXAMPLE 3 4-Vessel PSA Configuration using a Portion of Recycled Exhaust

FIGS. 5-6 show 4-vessel configurations that are illustrative but not limiting of another possible arrangement of adsorber vessels within the scope of the invention. In the configurations of FIGS. 5-6, at least one vessel is on feed at all times. Two or three of the other vessels can be undergoing other steps of the cycle, e.g., that regenerate the adsorbent, whilst the at least one vessel is making product at relatively high pressure. In the particular 4-vessel arrangement of FIG. 5, there is no fluid connection (i.e., exchange of gas streams) between the vessels, which reduces cycle complexity as well as equipment such as interconnecting piping and valving. However 4-vessel arrangements permit cycles to be conducted whereby two of the vessels not on feed can exchange gas, e.g., by employing a pressure equalization step whereby a first vessel at relatively higher pressure can be depressurized into a second vessel at relatively lower pressure, thereby increasing the pressure of the second vessel and decreasing the pressure of the first vessel, such as shown in FIG. 6. This type of cycle can permit higher (CO₂) recovery, since the stream(s) containing the target gas (CO₂) is(are) not released to the ambient atmosphere, but is(are) sent to another adsorbent vessel. The advantages of higher recovery from using a 4-vessel configurations (or configurations using more than 4 vessels) can be offset to some extent by the increased cost of added vessels, interconnecting piping and valving, and associated cycle control equipment required.

EXAMPLE 4 Integration of a HotPSA with Fuel Cell for Low Emissions Power Generation

FIG. 7 shows an example of a conventional integrated system where two molten carbonate fuel cells (MCFC) are combined with a natural gas combined cycle (NGCC). Relevant streams in FIG. 7 are described with their properties listed in Table 1 below.

TABLE 1 Thermodynamic properties and chemical compositions of the relevant streams in plant of FIG. 7 T (° C.) P (bar) M (kg/s) Ar CH₄ CO CO₂ H₂ H₂O N₂ O₂ C₂H₃ 1 15.0 1.0 656.7 0.9 — — 0.0 — 1.0 77.3 20.7 — 2 587.1 1.1 672.0 0.9 — — 3.9 — 8.3 74.4 12.5 — 3 750.0 1.0 681.0 0.9 — — 5.0 — 9.7 73.3 11.1 — 4 580.0 1.0 681.0 0.9 — — 5.0 — 9.7 73.3 11.1 — 5 660.0 1.0 642.1 0.9 — — 2.4 — 10.1 76.5 10.2 — 6 580.0 1.0 642.1 0.9 — — 2.4 — 10.1 76.5 10.2 — 7 617.2 1.0 628.8 0.9 — — 1.5 — 9.6 77.6 10.4 — 8 70.0 0.9 622.6 0.9 — — 1.5 — 9.6 77.6 10.4 — 9 15.0 70.0 21.6 — — — — — — — — — 10 200.0 69.0 15.3 — — — — — — — — — 11 200.0 69.0 1.3 — — — — — — — — — 12 600.0 24.0 9.0 — 14.4 17.1 24.0 42.5 — — — 1.3 13 15.0 70.0 5.1 — — — — — — — — — 14 600.0 1.1 24.2 — 18.6 — 0.4 — 79.1 0.2 — 1.7 15 732.2 1.1 24.2 — 0.1 8.2 7.3 51.5 32.6 0.1 — — 16 660.0 1.0 57.0 — 0.1 3.2 38.5 7.1 51.0 0.1 — — 17 640.0 1.0 6.2 — 0.1 8.2 7.3 51.5 32.6 0.1 — — 18 610.0 1.0 19.4 — 0.1 2.7 39.0 7.6 50.5 0.1 — — 19 65.0 1.0 76.4 — 0.1 3.1 36.6 7.2 50.9 0.1 — — 20 25.0 110.0 43.2 — <0.01 0.1 98.9 0.2 — 0.7 — — 21 100.0 24.5 7.6 — 0.5 20.0 28.1 50.4 — 0.6 0.3 — 22 340.0 130.0 48.1 — — — — — 100.0 — — — 23 559.0 119.6 135.0 — — — — — 100.0 — — — 24 562.0 27.6 140.6 — — — — — 100.0 — — — 25 351.7 6.9 144.1 — — — — — 100.0 — — — 26 150.8 1.1 19.2 — — — — — 100.0 — — — 27 233.8 30.0 11.5 — — — — — 100.0 — —

In the example of conventional technology, two MCFCs include a cathode with a cathode catalyst, an anode with an anode catalyst and an electrolyte placed between them. The reaction in the cathode requires oxygen (O₂) and carbon dioxide (CO₂) as reactants:

½O₂+CO₂+2e⁻→CO₃ ²⁻  (1)

The resulting product of the reaction includes carbonate ions CO₃ ²⁻, which are transported to the anode through the carbonate electrolyte. Further, the reaction in the anode requires the carbonate ions (CO₃ ²⁻) and hydrogen (H₂) as reactants:

CO₃ ²⁻+H₂→H₂O+CO₂+2e⁻  (2)

As illustrated in FIG. 7, stream 14 flows primarily CH₄ (18.6% in Table 1) and H₂O (79.1% in Table 1) to a reformer. The following reaction occurs in the reformer:

CH₄+2H₂O+heat→4H₂+CO₂   (3)

In other words, in the reformer, a portion of the natural gas introduced into the system reacts with water in an endothermic reaction and produces fuel gas (H₂) to be used in the anodes of the MCFCs and CO₂. In this example of the conventional technology, stream 17 includes the processed gas from the reformer to a heat exchanger and further to the anodes of the MCFCs at a mass flow of 6.2 kg/s, and at a temperature of 640° C. On the other hand, stream 4 feeds gas turbine exhaust gas into one of the cathodes of the MCFCs at a mass flow 681 kg/s with a significant concentration of oxygen (11.1%) and CO₂ (5%). Similarly, stream 6 feeds the gas turbine exhaust gas into the other one of the cathodes of the MCFCs at a mass flow 642.1 kg/s. The stream 6 includes 10.2% O₂ and 2.4% CO₂.

Further, the components of the streams 4 and 6 react according to equation (1) and produce CO₃ ²⁻ ions, which are transported to the anodes. Next, according to the current technology, H₂ produced in the process of reforming reacts with the CO₃ ²⁻ ions as indicated in equation (2), and electric potential difference is created. Thus, in addition to the generated electric energy, H₂O and CO₂ depart the anodes of the MCFCs in streams 16 and 18 as fuel cell exhaust gas. Of note, the exhaust gas of the streams 16 and 18 includes 7-8% of unreacted H₂. In the conventional technology, this causes significant energy inefficiency at the stage when CO₂ is disposed and stored.

Namely, as shown in FIG. 7, the streams 16 and 18 flow the anode exhaust gas at a high temperature (660° C.) into a heat recovery steam generator (HRSG). HRSG of the current technology cools the exhaust gas with a high degree of inefficiency. The cooling of the exhaust gas is necessary in order to lower the temperature down to 65° C. (stream 19) and prepare the mixture for the conditions required to adequately execute cryogenic compression.

Another step that renders storage of CO₂ expensive is the fact that CO₂ arrives at the cryogenic chamber mixed with H₂. For this reason, the separation of CO₂ is performed by cryogenic compression, which is a process of liquefaction of CO₂, during which pressure of CO₂ must reach a critical high value of 110 bara. According to the conventional technique presented in FIG. 7, the liquefied CO₂ is separated from gaseous H₂ and stored.

The requirement to cool down the exhaust gas from >600° C. to the ambient temperature prior to entering the cryogenic chamber and to compress the gas to 110 bara in the chamber produces a notable energy loss and adds a highly undesirable cost on the operation of the power plant system.

FIG. 8 is an example of the present invention that addresses the above discussed drawbacks of the existing methodology. In one embodiment, a high temperature PSA (HotPSA) is mounted at the exhaust of a MCFC and the stream 16, for example, may be fed into the HotPSA. As shown in Table 1, this may entail that an exhaust gas at a pressure of 1 bara and a temperature of 660° C. is fed into the HotPSA in order for the HotPSA to adsorb CO₂ at a high temperature and separate it from H₂. Thus, the adsorbed CO₂ may be harvested from the adsorbent by applying steam, in which case the CO₂ may be isolated simply by condensing the steam out of the mixture. This may eliminate the cost of an expensive and inefficient HRSG equipment, which has been conventionally used to cool down the CO₂ prior to cryogenic separation.

As a consequence, enriched H₂ may flow through the HotPSA as a light component of the adsorption, ready to be used as a reactant, for example, in another fuel cell. In addition, separating CO₂ from H₂ at high temperatures may also avoid the need to liquefy the CO₂ in order to separate it from the gaseous H₂. Moreover, the reused enriched H₂ may increase the fuel gas utilization of the power plant system.

Further, HotPSA may be purged with steam in order to harvest the adsorbed CO₂. The purged CO₂ may be recycled back to the stream 6 and into a cathode of a MCFC, for example, where it may react with O₂ and produce CO₃ ²⁻ ions, required for power generation of the MCFC. In this manner, yet another portion of the exhausted CO₂ would be captured, which would also decrease the demands of the CO₂ storage process and improve energy efficiency.

EXAMPLE 5 Integration of a Low Pressure Ratio HotPSA with Fuel Cell for Low Emissions Power Generation

A HotPSA may be in communication with the stream 16, which may flow the exhaust gas out of a MCFC at a pressure of 1 bara and a temperature of 660° C., for example. Thus, the incoming gas at such a low pressure may not produce a sufficient pressure difference to form an adsorption isotherm within the PSA required to manipulate loading of the heavy component within a reactor.

In order to address this potential issue, in the example in FIG. 9(a), a nitrogen jet ejector using N₂ as motive fluid may draw the MCFC anode gas containing, for example, CO₂ into a vessel containing HotPSA adsorbent. The addition of the ejector may facilitate the feeding of the low pressure MCFC exhaust gas into the HotPSA and increase the pressure across the adsorbent. The adsorbent material of the HotPSA may be capable of adsorbing CO₂ selectively from 350° C. to 650° C. Subsequent to this feed step, a low or ambient pressure steam purge may be employed to purge the vessel and produce enriched CO₂.

In the example in FIG. 9(b), steam may be used to create a suction effect on the far-side of the PSA vessel again employing an ejector, thus enabling a vacuum (i.e. low pressure ratio) swing cycle. Similar to the example in FIG. 9(a), the flow of the low pressure MCFC exhaust gas across the HotPSA adsorbent may be improved accordingly. The steam creating a vacuum may be used to purge another vessel in the PSA cycle.

EXAMPLE 6 Integration of a HotPSA with Ultra Low Emission Turbine for Low Emissions Power Generation

FIG. 10 illustrates an example of advanced power generation method of an Ultra-Low Emission Turbine (ULET) that may simultaneously produce power while reducing CO₂ emissions, sequestering CO₂ and producing inert N₂ from flue gases for subsequent industrial use.

ULET may improve utilization of excess air of the power generating system by recirculating O₂ produced in HRSG through one compressor and adding stoichiometric air through another compressor. In addition, the system may include a condenser that may condense the water produced in the process of hydrocarbon combustion and remove the water from the recirculation.

Thereby, ULET may remove oxygen and water from the exhaust gas, and desirably increase concentration of CO₂ to a level of, for example, 10% and higher. This concentration level may improve adsorption of CO₂ in comparison with lower concentrations of ˜4% produced by, for example, a natural gas fired combined cycles (NGCC) turbine. Moreover, ULET may also produce exhaust gas at a pressure as high as 19 bara, which is substantially higher than the pressure of 1 bara at the exhaust gas of the NGCC turbine. Consequently, the high pressure at the exhaust of ULET would improve both the adsorbing isotherm formation and the purging of the PSA.

However, the gas mixture at the exhaust of ULET, which may consist of CO₂ and N₂, for example, may exit ULET at a temperature as high as 420° C., for example. Thus, a HotPSA may be connected at the exhaust of ULET, in order to adsorb the high-temperature exhaust gas. In this manner, the HotPSA described above may be implemented to take advantage of the high pressure and the high concentration of CO₂ that ULET creates, while performing effective adsorption of CO₂ even at such high temperatures.

Although the present invention has been described in terms of specific embodiments, it need not necessarily be so limited. Suitable alterations/modifications for operation under specific conditions should be apparent to those skilled in the art. It is therefore intended that the following claims be interpreted as covering all such alterations/modifications as fall within the true spirit/scope of the invention. 

1. A method for optimizing emission reduction of a power generation system, comprising: passing hydrocarbons to a reformer for processing and producing H₂, passing the produced H₂ to a first anode of a first molten carbonate fuel cell, passing CO₂ and O₂ to a first cathode of the first molten carbonate fuel cell for processing and producing carbonate ions, transporting the produced carbonate ions from the first cathode to the first anode through a first electrolyte of the first molten carbonate fuel cell, generating power and an exhaust gas by reacting the H₂ with the transported carbonate ions in the first anode of the first molten carbonate fuel cell, the exhaust gas including CO₂ at a temperature greater than 600° C., passing the exhaust gas of the of the first molten carbonate fuel cell to a first swing adsorption reactor, and adsorbing the CO₂ from the exhaust gas on a first adsorbent material of the first swing adsorption reactor.
 2. The method of claim 1, further comprising: purging the first swing adsorption reactor with steam and removing the adsorbed CO₂.
 3. The method of claim 1, further comprising: recovering a H₂ stream unadsorbed by the first swing adsorption reactor.
 4. The method of claim 3, wherein the first swing adsorption reactor is between the first molten carbonate fuel cell and a second molten carbonate fuel cell.
 5. The method of claim 4, further comprising: passing the recovered H₂ stream unadsorbed by the first swing adsorption reactor to a second anode of the second molten carbonate fuel cell.
 6. The method of claim 5, further comprising: passing the CO₂ adsorbed by the first swing adsorption reactor to a second cathode of the second molten carbonate fuel cell.
 7. The method of claim 6, further comprising: passing the recovered H₂ stream unadsorbed by the first swing adsorption to the second anode of the second molten carbonate fuel cell, using the CO₂ adsorbed by the first swing adsorption reactor for producing carbonate ions in the second cathode of the second molten carbonate fuel cell, transporting the produced carbonate ions from the second cathode to the second anode through a second electrolyte of the second molten carbonate fuel cell, and generating power and an exhaust gas by reacting the recovered H₂ with the transported carbonate ions in the second anode of the second molten carbonate fuel cell.
 8. The method of claim 7, further comprising: passing the exhaust gas of the of the second molten carbonate fuel cell to a second swing adsorption reactor, wherein the exhaust gas includes CO₂ at a temperature greater than 600° C., and adsorbing the CO₂ from the exhaust gas on a second adsorbent material of the second swing adsorption reactor.
 9. The method of claim 1, wherein the step of passing the exhaust gas of the first molten carbonate fuel cell to the first swing adsorption reactor includes using an ejector to drive a motive fluid under a pressure that is higher than a pressure of the exhaust gas of the first molten carbonate fuel cell.
 10. The method of claim 9, wherein the ejector is placed at a far side of the first swing adsorption reactor with respect to the first molten carbonate fuel cell, and the ejector produces a vacuum inside the first swing adsorption reactor by driving the motive fluid.
 11. A method for optimizing emission reduction of a power generation system, comprising: passing hydrocarbons to a reformer for processing and producing H₂, passing the produced H₂ to a first anode of a first molten carbonate fuel cell, passing CO₂ and O₂ to a first cathode of the first molten carbonate fuel cell for processing and producing carbonate ions, transporting the produced carbonate ions from the first cathode to the first anode through a first electrolyte of the first molten carbonate fuel cell, generating power and an exhaust gas by reacting the H₂ with the transported carbonate ions in the first anode of the first molten carbonate fuel cell, the exhaust gas including CO₂ at a temperature greater than 600° C., passing the exhaust gas of the of the first molten carbonate fuel cell to a first swing adsorption reactor, and adsorbing the CO₂ from the exhaust gas on a first adsorbent material of the first swing adsorption reactor, wherein the first swing adsorption reactor is between the first molten carbonate fuel cell and a second molten carbonate fuel cell. 